CN102648169A - Method for converting mixed lower alkanes into aromatic hydrocarbons - Google Patents

Abstract

A process for the conversion of mixed lower alkanes into aromatics which comprises first reacting a mixed lower alkane feed comprising at least propane and ethane in the presence of an aromatization catalyst under reaction conditions which maximize the conversion of propane into first stage aromatic reaction products, separating ethane from the first stage aromatic reaction products, reacting ethane in the presence of an aromatization catalyst under reaction conditions which maximize the conversion of ethane into second stage aromatic reaction products, and optionally separating ethane from the second stage aromatic reaction products.

Description

Method for converting mixed lower alkanes into aromatic hydrocarbons

technical field

This invention relates to a process for the production of aromatic hydrocarbons from mixed lower alkanes. More specifically, the present invention relates to a two-stage process for increasing benzene production from a mixture of ethane and propane or ethane, propane and butane during dehydroaromatization.

Background technique

There is an expected global shortage of benzene, which is needed to make key petrochemicals such as styrene, phenol, nylon and polyurethane. Typically, benzene and other aromatics are separated from non-aromatics by using solvent extraction methods to separate aromatic-rich feedstock fractions, such as reformate from catalytic reforming and pyrolysis gasoline from naphtha cracking and obtained.

In the face of this anticipated supply shortfall, a number of catalysts and processes have been investigated for the production of aromatics, including benzene, from alkanes containing six or fewer carbon atoms per molecule. These catalysts are usually bifunctional, containing a zeolite or molecular sieve material to provide acidity, and one or more metals such as Pt, Ga, Zn, Mo, etc. to provide dehydrogenation activity. For example, US Pat. No. 4,350,835 describes the conversion of ethane-containing gaseous feeds to aromatic hydrocarbons using crystalline zeolite catalysts of the ZSM-5-type family containing small amounts of Ga. As another example, US Patent 7,186,871 describes the aromatization of C1-C4 alkanes using catalysts containing Pt and ZSM-5.

Most lower alkane dehydroaromatization processes use a one-step reaction. For example, EP 0147111 describes an aromatization process in which a C3-C4 feed is mixed with ethane and all reacted together in a single reactor. A few of these methods include two separate steps or stages. For example, US 3,827,968 describes a process involving oligomerization followed by aromatization. US 4,554,393 and US 4,861,932 describe a two-step process for propane comprising dehydrogenation followed by aromatization. None of these examples mention a two-stage process in which lower alkanes are aromatized in both stages.

The ease of conversion of individual alkanes to aromatics increases with increasing carbon number. When a mixed feed consisting of ethane and higher hydrocarbons is converted to benzene plus higher aromatics in a single stage, the selected reaction intensity is determined by the desired overall hydrocarbon conversion target. This may result in aggressively operating such a one-step process at higher temperatures if a significant level of ethane conversion is desired or required. A negative consequence of this higher intensity is non-selective side reactions of higher hydrocarbons such as propane leading to excessive hydrogenolysis to lower value methane. The end result is a significant reduction in the overall yield of benzene and other aromatics.

A process for the dehydroaromatization of light alkanes is provided wherein (a) the conversion of components in a mixed alkane feed can be optimized, (b) the final yield of benzene is higher than any other single aromatic product, and ( c) It would be advantageous to have a reduced generation of unwanted methane by-products.

Summary of the invention

According to the present invention, the above problems are solved by decoupling the optimum intensity for propane from the optimum intensity for ethane. This is achieved by devising a two-stage process as described below.

The present invention provides a process for the conversion of mixed lower alkanes to aromatics comprising first feeding a mixed lower alkane feed comprising at least propane and ethane in the presence of an aromatization catalyst at maximum propane (and optionally reacting under first stage reaction conditions in which any higher hydrocarbons in the feed, such as butane, are converted to a first stage aromatic reaction product which is mixed with unreacted ethane (and optionally Any other non-aromatic hydrocarbons that may be produced in the reaction, including ethane) are separated, and the ethane (and optionally at least a portion of any other non-aromatic hydrocarbons) is separated in the presence of an aromatization catalyst in a manner that maximizes ethane (and optionally in Any other non-aromatic hydrocarbons that may be produced in the first stage) are reacted under the reaction conditions of the second stage to convert the aromatic reaction products of the second stage, and optionally combine ethane (and any other non-aromatic hydrocarbons that may be produced) with The second stage is separation of aromatic reaction products.

In either or both of the first and second stages, a fuel gas consisting essentially of methane and hydrogen may also be produced. The fuel gas may be separated from the aromatic reaction products at either or both of the stages. Thus, fuel gas may be an additional product of the process of the present invention.

Description of drawings

Figure 1 is a schematic flow diagram illustrating a process diagram for the production of aromatics (benzene and higher aromatics) from a mixed lower alkane feed containing at least ethane and propane using a one-stage reactor-regenerator process.

Figure 2 is a schematic flow diagram illustrating a process diagram for the production of aromatics (benzene and higher aromatics) from a mixed lower alkane feed containing at least ethane and propane using a two-stage reactor-regenerator system.

Contents of the invention

The present invention is a process for the production of aromatic hydrocarbons, comprising at least 20 wt% ethane and at least 20 wt% A hydrocarbon feedstock of propane and possibly other hydrocarbons such as butane is contacted with a catalyst composition suitable for promoting the reaction of such hydrocarbons to aromatic hydrocarbons such as benzene. The gas hourly space velocity (GHSV) per hour may range from about 300 to about 6000. These conditions are used for each stage, but the conditions can be the same or different within a stage. The conditions can be optimized for the conversion of propane, and possibly other higher alkanes such as butane, in the first stage, and ethane in the second stage. In the first stage, the reaction temperature is preferably in the range of about 400 to about 650°C, most preferably in the range of about 420 to about 650°C, and in the second stage, the reaction temperature is preferably in the range of about 450 to about 680°C, most preferably About 450 to about 660°C. The main desired products of the process of the present invention are benzene, toluene and xylenes (BTX). In one embodiment, the first stage reaction conditions can be optimized for the conversion of propane to BTX. Optionally, the first stage reaction conditions can also be optimized for the conversion of any higher hydrocarbons that may be present in the feedstock to BTX. In another embodiment, the second stage reaction conditions can be optimized for the conversion of ethane to BTX. Optionally, the second stage reaction conditions can also be optimized for the conversion of any other non-aromatic hydrocarbons that may be produced in the first stage to BTX.

The first stage and second stage reactors can be operated under similar conditions. When either reactor is operated at a higher temperature, ie above about 630-650°C, more fuel gas and less aromatics are produced even though the net feed per pass conversion for this stage may be higher. Therefore, it is better to operate at lower temperatures and convert less feed in each pass of the stages in order to produce more aromatics overall, even though more ethane must be recycled. Operating within the preferred range helps maximize aromatics production by minimizing fuel gas production. Use of higher temperatures may maximize fuel gas production.

Fuel gas may be an additional product of the process of the invention. The fuel gas mainly includes methane and hydrogen produced together with aromatic hydrocarbons. The fuel gas can be used to generate electricity and/or generate steam. The hydrogen in the fuel gas can be separated and used in refining or chemical reactions requiring hydrogen, including the hydrodealkylation of toluene and/or xylenes discussed below.

It is possible to carry out the process in batch mode, using separate reactors for each stage or using the same reactor for each stage, but it is highly preferred to carry out the process in continuous mode in separate reactors. Each stage can be carried out in a single reactor or in two or more reactors connected in parallel. Preferably, at least two reactors are used per stage, so that one reactor can be used for aromatization while the other is taken offline so that the catalyst can be regenerated. The aromatization reactor system can be a fluidized bed, moving bed or circulating fixed bed design. A circulating fixed bed design is preferred for use in the present invention.

The hydrocarbons in the feedstock may comprise at least about 20 wt% propane, at least about 20 wt% ethane, and optionally at least about 10 to 20 wt% butane, pentane, and the like. In one embodiment, the feedstock is about 30 to about 50 wt% propane and about 30 to about 50 wt% ethane. The feed may contain small amounts of _C2 - _C4 olefins, preferably no more than 5 to 10 weight percent. Too much olefin may result in unacceptable amounts of coking and deactivation of the catalyst.

The mixed propane/ethane or mixed C2-C4 lower alkane feed stream can be derived, for example, from an ethane/propane rich stream derived from natural gas, refinery or petrochemical streams including waste streams. Examples of potentially suitable feed streams include, but are not limited to, residual ethane and propane and butane from natural gas (methane) purification, pure ethane and propane and butane co-produced at liquefied natural gas (LNG) sites streams (also known as natural gas liquids), C2-C4 streams from co-produced associated gases from crude oil production (they are usually too small to build an LNG plant, but may be sufficient to build a chemical plant), unreacted streams from steam crackers The "scrap" stream from the naphtha reformer, and the C1-C4 by-product stream from the naphtha reformer (the latter two are of low value in some markets such as the Middle East).

Typically, natural gas, mainly comprising methane, enters an LNG plant at elevated pressure and is pretreated to produce a purified feed suitable for liquefaction at cryogenic temperatures. Separation of ethane, propane, butane and other gases from methane. The purified gas (methane) is processed through multiple cooling stages using heat exchangers to gradually reduce its temperature until liquefaction is achieved. The separated gas can be used as a feed stream to the present invention. The by-product stream produced by the process of the present invention may have to be cooled for storage or recycle, which can be accomplished using heat exchangers that cool the purified methane gas.

Any of a variety of catalysts can be used to facilitate the reaction of propane and ethane and possibly other alkanes to aromatics. One such catalyst is described in U.S. 4,899,006, which is hereby incorporated by reference in its entirety. The catalyst compositions described therein comprise aluminosilicates on which gallium is deposited and/or aluminosilicates in which cations have been exchanged for gallium ions. The molar ratio of silica to alumina is at least 5:1.

Another catalyst that can be used in the process of the invention is described in EP 0244162. This catalyst comprises the catalyst described in the preceding paragraph and rhodium and platinum selected from group VIII metals. The aluminosilicate is said to be preferably of MFI or MEL type structure and may be ZSM-5, ZSM-8, ZSM-11, ZSM-12 or ZSM-35.

Other catalysts that may be used in the process of the present invention are described in U.S. 7,186,871 and U.S. 7,186,872, both of which are hereby incorporated by reference in their entirety. The first patent describes a platinum-containing ZSM-5 crystalline zeolite synthesized by preparing a zeolite containing aluminum and silicon in its structure, precipitating platinum on the zeolite, and calcining the zeolite. The second patent describes a catalyst that includes gallium in its structure and is substantially free of aluminum.

Preferably, the catalyst consists of a zeolite, a platinum group noble metal that promotes the dehydrogenation reaction, and a second inert or less reactive metal that will attenuate the C2 and higher hydrocarbons in the feed that the noble metal will feed Propensity for catalytic hydrogenolysis to methane and/or ethane. Attenuating metals that can be used include those described below.

Other catalysts that may be used in the process of the present invention include those described in U.S. 5,227,557, which is hereby incorporated by reference in its entirety. These catalysts comprise MFI zeolite together with at least one platinum group noble metal and at least one other metal selected from tin, germanium, lead and indium.

One preferred catalyst for use in the present invention is described in US Application No. 12/371787, filed February 16, 2009, entitled "Process for the Conversion of Ethane to Aromatic Hydrocarbons." This application is hereby incorporated by reference in its entirety. This application describes a catalyst comprising: (1) 0.005 to 0.1 wt% (weight %), on a metal basis, platinum, preferably 0.01 to 0.05 wt%, (2) on a metal basis, selected from the group consisting of tin, The amount of lead and germanium impairing metals is preferably not more than 0.2 wt% of the catalyst, and wherein the amount of platinum can be more than the amount of impairing metals not exceeding 0.02 wt%; (3) 10 to 99.9 wt% aluminosilicate , preferably zeolite, based on aluminosilicate, preferably 30 to 99.9 wt%, preferably selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23 or ZSM-35, preferably converted to the H+ form, preferably SiO ₂ /Al ₂ O ₃ molar ratio of 20:1 to 80:1, and (4) a binder, preferably selected from silica, alumina and mixtures thereof.

Another preferred catalyst for use in the present invention is described in US Provisional Application No. 61/029939, filed February 20, 2008, entitled "Process for the Conversion of Ethane to Aromatic Hydrocarbons." This application is hereby incorporated by reference in its entirety. Said application describes a catalyst comprising: (1) 0.005 to 0.1 wt % (weight %) platinum, preferably 0.01 to 0.06 wt %, most preferably 0.01 to 0.05 wt %, on a metal basis, (2) The amount of iron is equal to or greater than the amount of platinum, but based on the metal, not exceeding 0.50 wt%, preferably not exceeding 0.20 wt%, most preferably not exceeding 0.10 wt% of the catalyst; (3) 10 to 99.9 wt% % aluminosilicate, preferably zeolite, preferably 30 to 99.9 wt%, based on aluminosilicate, preferably selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23 or ZSM-35, preferably converted in the H+ form, preferably in a SiO ₂ /Al ₂ O ₃ molar ratio of 20:1 to 80:1, and (4) a binder, preferably selected from silica, alumina and mixtures thereof.

Another preferred catalyst for use in the present invention is described in US Application No. 12/371803, filed February 16, 2009, entitled "Process for the Conversion of Ethane to Aromatic Hydrocarbons." This application is hereby incorporated by reference in its entirety. This application describes a catalyst comprising: (1) 0.005 to 0.1 wt% (weight %) platinum, preferably 0.01 to 0.05 wt%, most preferably 0.02 to 0.05 wt%, on a metal basis, (2) gallium The amount is equal to or greater than the amount of platinum, based on the metal, preferably not more than 1 wt%, most preferably not more than 0.5 wt%; (3) 10 to 99.9 wt% of aluminosilicate, preferably zeolite, in the form of aluminosilicate As a basis, preferably 30 to 99.9 wt%, preferably selected from ZSM-5, ZSM-11, ZSM-12, ZSM-23 or ZSM-35, preferably converted to H+ form, preferably SiO ₂ /Al ₂ O ₃ molar ratio is 20:1 to 80:1, and (4) a binder, preferably selected from silica, alumina and mixtures thereof.

One undesirable product in aromatization reactions is coke, which can deactivate the catalyst. While the catalyst and operating conditions and reactors are selected to minimize coke production, it is often necessary to regenerate the catalyst at times during its useful life. Regeneration can increase the useful life of the catalyst.

Regeneration of coked catalysts has been practiced commercially for decades, and various regeneration methods are known to those skilled in the art.

Regeneration of the catalyst can be carried out in the aromatization reactor or in a separate regeneration vessel or reactor. For example, the catalyst can be regenerated by firing the coke in the presence of an oxygen-containing gas at elevated temperatures as described in US Patent No. 4,795,845, which is hereby incorporated by reference in its entirety. Regeneration with air and nitrogen is illustrated in the Examples in US Patent No. 4,613,716, which is hereby incorporated by reference in its entirety. Other possible methods include air calcination, hydrogen reduction, and treatment with sulfur or sulfide materials. Platinum catalysts have been used to assist in the combustion of coke deposited on such catalysts.

Preferred regeneration temperatures for use herein range from about 450 to about 788°C. The preferred temperature range for regeneration in the first stage is from about 470 to about 788°C. The preferred temperature range for regeneration in the second stage is from about 500 to about 788°C.

Unreacted methane and by-product hydrocarbons can be used in other steps, stored and/or recycled. It may be necessary to cool these by-products to liquefy them. When ethane or mixed lower alkanes originate from an LNG plant as a result of natural gas purification, at least some of the by-products can be cooled and liquefied using heat exchangers used to liquefy purified natural gas (methane).

Toluene and xylenes can be converted to benzene by hydrodealkylation. Hydrodealkylation involves the reaction of toluene, xylene, ethylbenzene, and higher aromatics with hydrogen to strip alkyl groups from aromatic rings, producing additional benzene and light ends including methane and ethane separated from benzene. This step significantly increases the overall yield of benzene and is therefore highly advantageous.

Both thermal and catalytic hydrodealkylation processes are known in the art. A hydrodealkylation process is described in US Published Patent Application No. 2009/0156870, which is hereby incorporated by reference in its entirety.

The overall process of the present invention may also include the reaction of benzene with propylene to form cumene, which may in turn be converted to phenol and/or acetone. Propylene may be produced separately in a propane dehydrogenation unit, or may come from an effluent stream of an olefin cracking process or other sources. A process for reacting benzene with propylene to produce cumene is described in US Published Patent Application No. 2009/0156870, which is hereby incorporated by reference in its entirety.

The overall process of the present invention may also include the reaction of benzene with an olefin such as ethylene. Ethylene can be produced separately in an ethane dehydrogenation unit, or can come from an effluent stream of an olefin cracking process or other sources. Ethylbenzene is an organic compound and an aromatic hydrocarbon. Its main use is in the petrochemical industry as an intermediate compound in the production of styrene, which in turn is used to make polystyrene, a commonly used plastic material. A process for the production of ethylbenzene by reacting benzene with ethylene is described in US Published Patent Application No. 2009/0156870, which is hereby incorporated by reference in its entirety.

Styrene can then be produced by dehydrogenation of ethylbenzene. One method of producing styrene is described in US Patent No. 4,857,498, which is hereby incorporated by reference in its entirety. Another method of producing styrene is described in US Patent No. 7,276,636, which is hereby incorporated by reference in its entirety.

Example

The following examples are provided for illustrative purposes only and do not limit the scope of the invention.

Example 1

In this example, laboratory test results were used to illustrate a one-stage aromatization process compared to a two-stage process using the same catalyst in each stage. The lower alkane feed in this example consisted of 50 wt% each of ethane and propane, and the temperature of the second stage was higher than that of the first stage.

Catalyst A was produced on cylindrical extruded particles with a diameter of 1.6 mm containing 80 wt% powder of zeolite ZSM- _5CBV 2314 ( _SiO2 / _Al2O3 molar ratio 23:1, available from Zeolyst International) and 20 wt% alumina binder. The extrudate samples were calcined in air up to 650°C to remove residual moisture prior to use in catalyst preparation. Catalyst A had a target metal loading of 0.025%w Pt and 0.09wt% Ga.

By first combining appropriate aqueous solutions of tetraammineplatinum nitrate and gallium(III) nitrate, diluting the mixture with deionized water to a volume just sufficient to fill the pores of the above ZSM-5/alumina extrudate, and The extrudates were impregnated with this solution at atmospheric pressure, thereby depositing metal on 25-100 grams of the extrudate samples. After dipping, the samples were aged at room temperature for 2-3 hours and then dried overnight at 100°C.

Catalyst A was subjected to the three performance tests described below. Performance Test 1 was conducted under conditions that can be used for a one-stage aromatization process using a mixed ethane/propane feed. Performance tests 2 and 3 were carried out under conditions applicable to the first and second stages, respectively, of the two-stage aromatization process of the present invention.

For each of the three performance tests, 15-cc of fresh (previously untested) catalyst charge was loaded "as is" without crushing into Type 316H stainless steel tubing (1.40 cm inside diameter) and placed in the same In a four-zone furnace connected by a gas flow system.

Before the performance test 1, the newly installed catalyst A was pretreated in situ as follows under atmospheric pressure (about 0.1 MPa absolute pressure):

(a) Calcination with approximately 60 liters per hour (L/hr) of air, during which the reactor wall temperature increased from 25°C to 510°C within 12 hours, maintained at 510°C for 4-8 hours, and then at 1 Further increase from 510°C to 630°C within 1 hour, and then keep at 630°C for 30 minutes;

(b) Purging with nitrogen gas at about 60L/hr, 630°C for 20min;

(c) Reduction with 60 L/hr of hydrogen for 30 min, during which the temperature of the reactor wall increased from 630°C to 675°C.

At the end of the above reduction step, the hydrogen flow was terminated, and the catalyst was fed at atmospheric pressure (approximately 0.1 MPa absolute pressure), 675° C. ), exposed to a feed consisting of 50 wt% ethane and 50 wt% propane. Three minutes after feed introduction, the total reactor outlet stream was sampled for analysis by an on-line gas chromatograph. From the composition data obtained from the gas chromatographic analysis, the initial ethane, propane and total conversions were calculated according to the following formula:

Ethane conversion, %=100×(wt% ethane in the feed-wt% ethane in the outlet stream)/(wt% ethane in the feed)

Propane conversion, % = 100 x (wt% propane in the feed - wt% propane in the outlet stream)/(wt% propane in the feed)

Total conversion of ethane+propane = ((wt% ethane in feed ×% ethane conversion) + (wt% propane in feed ×% propane conversion))/100

Performance Test 2 was performed in the same manner and under the same conditions as Performance Test 1 above, except that the final temperature during the air calcination pretreatment step was 600°C, the nitrogen purge and hydrogen reduction steps were performed at 600°C, and the ethane/ The propane feed was introduced at a reactor wall temperature of 600°C. This simulates the first stage of a two-stage process.

For performance test 3, the newly installed catalyst A was pretreated in situ as follows under atmospheric pressure (about 0.1 MPa absolute pressure):

(a) Calcination with approximately 60 liters per hour (L/hr) of air during which the reactor wall temperature is increased from 25°C to 510°C over 12 hours and then maintained at 510°C for 4-8 hours;

(b) About 60L/hr, 510°C nitrogen purging for 30min;

(c) Reduction with hydrogen at 60 L/hr for 2 hours.

At the end of the above reduction step, the hydrogen flow was terminated, and the catalyst was charged at atmospheric pressure (approximately 0.1 MPa absolute pressure), 510° C. ), exposed to a feed consisting of 100 wt% ethane. After 10 min at these conditions, the reactor wall temperature increased to 621°C. 25 min after the introduction of the ethane feed, the total reactor outlet stream was sampled by on-line gas chromatography for analysis. From the compositional data obtained from the gas chromatographic analysis, the initial ethane, propane and total conversions were calculated according to the formula given above:

Table 1 presents the results of on-line gas chromatographic analysis of the total product streams from Performance Tests 1-3 above.

Table 1

Performance test 1 2 3 Catalyst A A A Catalyst volume, cc 15 15 15 Reactor wall temperature, ℃ 675 600 621 Pressure, MPa 0.1 0.1 0.1 Feed composition Ethane, wt% 50 50 100 Propane, wt% 50 50 -0- Gross Feed Rate, GHSV 1000 1000 1000 Total Feed Rate, WHSV 1.93 1.88 1.61 Ethane conversion rate, % 51.50 1.10 49.28 Propane conversion rate, % 99.50 98.62 - Total conversion rate of ethane+propane, % 75.49 49.84 49.28 Reactor discharge composition, wt% hydrogen 5.43 3.44 4.71 methane 17.33 9.91 7.56 Vinyl 5.51 2.81 3.95 Ethane 24.26 49.47 50.72 Propylene 0.59 0.42 0.58 propane 0.25 0.69 0.70 C4 0.09 0.08 0.11 C5 -0- -0- -0- Benzene 26.64 18.40 16.60 Toluene 10.21 11.77 8.72 C8 aromatic hydrocarbons 1.47 2.67 1.70 C9+ aromatic hydrocarbons 8.22 0.34 4.65

Total Aromatic Hydrocarbons 46.54 33.18 31.67

As can be seen from Table 1, the one-stage process produces 46.54 wt% total aromatics from a given ethane/propane feed, while the two-stage process is based on 100 wt% total feed into stage 1, and then unconverted in stage 1 Ethane was fed to stage 2, producing 48.85 wt% total aromatics. In true two-stage operation, it is also possible that the feed to stage 2 includes all non-aromatic hydrocarbons in the effluent from stage 1 except for the fuel gases (methane and hydrogen). These non-aromatics would include not only unconverted ethane but also ethylene, propylene, propane, etc., based on 100 wt% total feed to stage 1, which would make it possible to increase the total aromatics yield to slightly over 50 wt %.

Example 2:

Method configuration comparison

2.1 One-stage method (comparative)

Figure 1 is a schematic flow diagram illustrating a process diagram for the production of aromatics (benzene and higher aromatics) from a feed containing 50 wt% ethane and 50 wt% propane using a one-stage reactor-regenerator process.

25 tons/hr (tph) stream 1 consisting mainly of 50/50 (ethane/propane) mixed feed (including small amounts of methane, butane, etc.), with mainly ethane and possibly including but not limited to ethylene, propylene , methane, other hydrocarbons of butane and some hydrogen are mixed in recycle stream 2. At this point the total feed stream 3 is introduced to the single stage aromatization reactor 100 . The aromatization reactor system can be a fluidized bed, moving bed or circulating fixed bed design. Here a circulating fixed bed design is used. The reactor system used "Catalyst A" as described in Example 1 above. Unconverted reactants as well as products exit reactor 100 via stream 4 and are sent to a separation system. Unconverted reactants and light hydrocarbons are recycled back to reactor 100 in stream 2, while the separation system yields fuel gases (mainly methane and hydrogen in stream 8 from vapor-liquid separator 200), C9 ₊ Liquid products as well as benzene, toluene and xylenes (BTX).

The reactor 100 operates at about 1 atmosphere pressure and at a temperature of 675°C, while the regenerator 300 operates at around 730°C, where the coke formed in the reactor 100 is removed. By preheating the heat during the regeneration step The catalyst solids mixture provides the heat required for the reaction step (9). In the regeneration step, the coke-containing catalyst flows through stream 5 to regenerator 300 and is supplied with stripping gas. The regenerated catalyst flows back to the reactor 100 via stream 6 and the stripping gas exits the regenerator 300 via stream 7 .

Reactor 100 achieved almost complete conversion of propane (greater than 99%) while converting about half of the ethane, as was the case in Table 1 and in Performance Test 1 above. The average conversion rate of single-pass mixed feed is 75.49%. As shown in Figure 1, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 2 below. This one-stage operation produces about 8.33 tph of benzene (from column 400 via stream 10), 3.25 tph of toluene (from column 500 via stream 11), and 0.5 tph of mixed xylenes (from column 500 via stream 12). 600), resulting in a total BTX yield of 49 wt% and a total liquid yield of 60 wt% relative to the mixed feed. Fuel gas generation (stream 8) was 9.2 tph which was 36.7 wt% of the mixed feed.

2.2 Two-stage approach

Figure 2 is a schematic flow diagram for the production of aromatics (benzene and higher aromatics) from a feed containing 50 wt% ethane and 50 wt% propane using the two-stage reactor-regenerator system of the present invention.

A mixed feed of 25 tons/hr (tph) primarily 50/50 ethane/propane, containing small amounts of methane, butane, etc. (stream 1) was fed to stage 1 using "Catalyst A" as described in Example 1 Aromatization reactor 100. The first stage reactor 100 operates at about 1 atmosphere pressure and at a temperature of about 600°C, while the stage 1 regenerator 200, which removes the coke formed in said reactor 100, operates at about 730°C. The heat required for the reaction step is provided by the preheated hot catalyst solids mixture during the regeneration step. Reactor 100 achieved almost complete conversion of propane (greater than 98%) with very little conversion of ethane, as was the case in Performance Test 2 of Table 2. Reactor effluent stream 3a is then mixed with the reactor effluent (stream 3b) of the second stage reactor 300 described below. The combined effluent from the two stage reactors (stream 4) is then fed to a separation system where the unconverted reactants consist primarily of ethane and some other hydrocarbons, possibly including ethylene, propane, propylene, methane, butane and light hydrocarbons along with some hydrogen are used as feed (stream 2) to Stage-2 aromatization reactor 300 using "Catalyst A" described above.

The second stage reactor 300 operates at about 1 atmosphere and a temperature of about 620°C, while the regenerator 400 operates at about 730°C, which removes coke formed in the reactor. The heat required for the reaction step is provided by the hot catalyst solids mixture which is preheated during the regeneration step. The second stage reactor 300 converted almost half of the ethane fed to it, as was the case for performance test 3 in Table 1 above. The effluent from the second stage reactor 300 (stream 3b) is mixed with the effluent from the first stage reactor 100 described above. Both Stage 1 and Stage 2 of the aromatization reactor system use a circulating fixed bed design.

The average single pass conversion of the mixed feed was obtained from the cumulative conversions of ethane and propane (feed) in the two stages and was calculated to be 74.5%. As shown in Figure 2, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 3 below. This two-stage operation produces about 8.6 tph of benzene (from column 600 via stream 10), 5.1 tph of toluene (from column 700 via stream 11), and 1.1 tph of xylenes (from column 800 via stream 12), compared to From the mixed feed, the resulting total BTX yield was 60 wt% and the total liquid yield was 65 wt%. The resulting undesired fuel gas (stream 8 from vapor-liquid separator 500) was about 7 tph, which was about 28 wt% of the mixed feed.

2.3 Comparison of method configurations

Table 2 below shows a comparison of the system performance of the one-stage and two-stage methods. Conditions of constant total feed conversion resulting from the process were compared. Compared to each of the two stages in a two-stage process, a single-stage process must be operated at higher temperatures to achieve similar feed conversions per pass. It is evident from Table 2 that the two-stage operation produces higher product yields of benzene, toluene, mixed xylenes and C9+ liquids with less undesirable fuel gas formation compared to the one-stage process.

Table 2

Notice:

• All yields are expressed as tons of product per ton of mixed feed entering the overall process, expressed as a percentage.

The average conversion rate per pass of the two-stage method is calculated as follows:

Total propane conversion x mole fraction of propane in mixed feed) + (total ethane conversion x mole fraction of ethane in mixed feed)

Example 3

In this example, the results of laboratory tests are used to illustrate a one-stage aromatization process compared to a two-stage process using a different catalyst in each stage. The lower alkane raw material in this embodiment is composed of 50 wt% each of ethane and propane, and the temperature of the second stage is higher than that of the first stage.

Catalyst B was prepared on zeolite ZSM-53024E powder (SiO ₂ /Al ₂ O ₃ molar ratio 30:1 ) available from Zeolyst International. ZSM-5 powder samples were calcined in air up to 650 °C to remove residual moisture before being used in catalyst preparation. The target metal loading for Catalyst B was 0.04 wt% each of platinum and tin.

By first combining appropriate stock solutions of tetraammineplatinum nitrate and tin(IV) nitrate in water, diluting the mixture with deionized water to a pore volume just sufficient to fill the powder, and impregnating the powder with this solution at room temperature and atmospheric pressure. The powder described above was used to deposit the metal on 50 grams of the ZSM-5 sample described above. The impregnated samples were aged at room temperature for about 2 hours and then dried overnight at 100°C.

The dried catalyst powder was packed into a plastic bag and pressed using an isostatic press at a pressure of 1380 barg for 2 min. The resulting "rocks" were then crushed and sieved to obtain 2.5-8.5 mm particles suitable for performance testing.

For Performance Test 4, 15-cc of fresh (previously untested), pressed and crushed Catalyst B was packed into Type 316H stainless steel tubing (1.40 cm inside diameter) and placed in a four-zone furnace connected to the airflow system . Performance test 4 was performed in the same manner and under the same conditions as performance test 3 in Example 1 above. Based on the compositional data obtained by gas chromatographic analysis, the initial ethane, propane and total conversions were calculated according to the formula given in Example 1 above.

Table 3 lists the results of the on-line gas chromatographic analysis of the total product streams from Performance Tests 1 and 2 (from Example 1) and Performance Test 4. Performance Test 1 was conducted under conditions for a one-stage aromatization process using a mixed ethane/propane feed. Performance tests 2 and 4 were carried out under conditions which can be used in the first and second stages, respectively, of the two-stage aromatization process of the present invention.

table 3

Performance test 1 2 4 Catalyst A A B Catalyst volume, cc 15 15 15 Reactor wall temperature, ℃ 675 600 621 Pressure, MPa 0.1 0.1 0.1 Feed composition Ethane, wt% 50 50 100

Propane, wt% 50 50 -0- Gross Feed Rate, GHSV 1000 1000 1000 Total Feed Rate, WHSV 1.93 1.88 2.05 Ethane conversion rate, % 51.50 1.10 50.97 Propane conversion rate, % 99.50 98.62 - Total conversion rate of ethane+propane, % 75.49 49.84 50.97 Reactor discharge composition, wt% hydrogen 5.43 3.44 4.78 Methane 17.33 9.91 8.48 Vinyl 5.51 2.81 4.23 Ethane 24.26 49.47 49.03 Propylene 0.59 0.42 0.54 propane 0.25 0.69 0.60 C4 0.09 0.08 0.10 C5 -0- -0- 0.01 Benzene 26.64 18.40 15.19 Toluene 10.21 11.77 8.93 C8 aromatic hydrocarbons 1.47 2.67 1.93 C9+ aromatic hydrocarbons 8.22 0.34 6.19 Total Aromatic Hydrocarbons 46.54 33.18 32.24

As can be seen from Table 3, the one-stage process produces 46.54 wt% total aromatics from a given ethane/propane feedstock, while the two-stage process is based on 100 wt% total feed to Alkanes were fed to stage 2, producing 49.13 wt% total aromatics. In true two-stage operation, it is also possible that the feed to stage 2 includes all non-aromatic hydrocarbons in the output of stage 1 except for the fuel gas (methane and hydrogen). These non-aromatics would include not only unconverted ethane but also ethylene, propylene, propane, etc., based on 100 wt% total feed to stage 1, which would make it possible to increase the total aromatics yield to slightly over 50 wt %.

Example 4

Method configuration comparison

4.1 One-stage method (comparative)

Figure 1 is a schematic flow diagram illustrating a process diagram for the production of aromatics (benzene and higher aromatics) from a feedstock containing 50 wt% ethane and 50 wt% propane using a reactor-regenerator one-stage process.

25 tons/hr (tph) stream 1 mainly composed of 50/50 (ethane/propane) mixed feedstock (including a small amount of methane, butane, etc.), and mainly composed of ethane and may include but not limited to ethylene, propylene , methane, other hydrocarbons of butane and some hydrogen are mixed in recycle stream 2. At this point the total feed stream 3 is introduced to the single stage aromatization reactor 100 . The aromatization reactor system can be a fluidized bed, moving bed or circulating fixed bed design. Here a circulating fixed bed design is used. The reactor system used "Catalyst A" as described in Example 1 above. Unconverted reactants as well as products exit reactor 100 via stream 4 and are fed to the separation system. Unconverted reactants and light hydrocarbons are recycled back to reactor 100 in stream 2, while the separation system yields fuel gases (mainly methane and hydrogen in stream 8 from vapor-liquid separator 200), C9+ liquid products and benzene, toluene, and xylenes (BTX).

Reactor 100 operates at about 1 atmosphere of pressure and at a temperature of 675° C., while regenerator 300 operates at around 730° C., in which coke formed in reactor 100 is removed. The heat required for the reaction step is provided by the preheated hot catalyst solids mixture during the regeneration step (9). In the regeneration step, the coke-containing catalyst flows through stream 5 to regenerator 300 and is supplied with stripping gas. The regenerated catalyst flows back to the reactor 100 via stream 6 and the stripping gas exits the regenerator 300 via stream 7 .

Reactor 100 achieved nearly complete conversion of propane (greater than 99%) while converting about half of the ethane, as in Table 1 and Performance Test 1 above. The average conversion rate of single-pass mixed feed is 75.49%. As shown in Figure 1, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 2 below. This one-stage mode of operation produces about 8.33 tph of benzene (from column 400 via stream 10), 3.25 tph of toluene (from column 500 via stream 11) and 0.5 tph of mixed xylenes (from column 500 via stream 12). 600), resulting in a total BTX yield of 49 wt% and a total liquid yield of 60 wt% relative to the mixed feed. The fuel gas produced (stream 8) was 9.2 tph which was 36.7 wt% of the mixed feed.

4.2 Two-stage approach

Figure 2 is a schematic flow diagram for the production of aromatics (benzene and higher aromatics) from a feed containing 50 wt% ethane and 50 wt% propane using the two-stage reactor-regenerator system of the present invention.

A 26 ton/hr (tph) mixed feed of primarily 50/50 ethane/propane with minor amounts of methane, butane, etc. (stream 1) was fed to stage 1 using "catalyst A" as described in Example 3 Aromatization reactor 100. The first stage reactor 100 operates at about 1 atmosphere pressure and at a temperature of about 600°C, while the stage 1 regenerator 200, which removes the coke formed in the reactor 100, operates at about 730°C. The heat required for the reaction step is provided by the hot catalyst solids mixture which is preheated during the regeneration step. Reactor 100 achieved almost complete conversion of propane (greater than 98%) with very little conversion of ethane, as was the case for Performance Test 2 in Table 3. Reactor effluent stream 3a is then mixed with the reactor effluent (stream 3b) of the second stage reactor 300 described below. The combined effluent from the two stage reactors (stream 4) is then fed to a separation system where the unconverted reactants consist primarily of ethane and some other hydrocarbons which may include ethylene, propane, propylene, methane, butane and light hydrocarbons along with some hydrogen are used as the feed (stream 2) to stage-2 aromatization reactor 300, which uses "catalyst B" as described above.

The second stage reactor 300 operates at about 1 atmosphere and a temperature of about 620°C, while the regenerator 400 operates at about 730°C, which removes coke formed in the reactor. The heat required for the reaction step is provided by the hot catalyst solids mixture which is preheated during the regeneration step. The second stage reactor 300 converted almost half of the ethane fed to it, as was the case for Performance Test 4 in Table 3 above. The effluent from the second stage reactor 300 (stream 3b) is mixed with the effluent from the first stage reactor 100 described above. Both Stage-1 and Stage-2 of the aromatization reactor system use a circulating fixed bed design.

The average single pass conversion of the mixed feed was obtained from the cumulative conversions of ethane and propane (feed) in the two stages and was calculated to be 75.35%. As shown in Figure 2, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 4 below. This two-stage operation produces about 8.4 tph of benzene (from column 600 via stream 10), 5.3 tph of toluene (from column 700 via stream 11), and 1.2 tph of xylenes (from column 800 via stream 12), compared to From the mixed feed, the resulting total BTX yield was 60 wt% and the total liquid yield was 57.3 wt%. Undesired fuel gas generation (stream 8 from vapor-liquid separator 500) was about 7.6 tph, which was about 29.2 wt% of the mixed feed.

4.3 Comparison of Method Configurations

Table 4 below shows a comparison of the system performance of the one-stage and two-stage approaches. Conditions under which the process yielded a constant overall feed conversion were compared. Compared to each of the two stages in a two-stage process, a single-stage process must be operated at higher temperatures to achieve similar feed conversions per pass. It is evident from Table 4 that the two-stage operation produced higher product yields of benzene, toluene, mixed xylenes and C9+ liquids with less undesirable fuel gas formation compared to the one-stage process.

Table 4

Notice:

All yields are expressed as tons of product per ton of mixed feed fed to the overall process, expressed as a percentage.

The average conversion per pass of the two-stage process was calculated as follows:

(total propane conversion x mole fraction of propane in mixed feed) + (total ethane conversion x mole fraction of ethane in mixed feed)

Example 5

In this example, the results of laboratory tests are used to illustrate a one-stage aromatization process compared to a two-stage process utilizing the same catalyst in each stage. The lower alkane raw material in this example is composed of 31.6wt% ethane, 29.5wt% propane and 38.9wt% n-butane, and the temperature of the second stage is higher than that of the first stage.

A fresh 15-cc charge of Catalyst A (prepared as described above in Example 1) was subjected to performance testing as described below. Performance Test 5 was conducted under conditions for a one-stage aromatization process using a mixed ethane/propane feed. Performance Test 6 was conducted under the conditions used in the first stage of the two-stage aromatization process of the present invention using a mixed ethane/propane/butane feed. Performance Test 3 (described in Example 1) was performed under the conditions used for the second stage of the two-stage aromatization process of the present invention.

Performance Test 5 was performed in the same manner and under the same conditions as used in Performance Test 1 (described in Example 1), except that the feed to Performance Test 5 consisted of 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% normal butane composition. Performance Test 6 was performed in the same manner and under the same conditions as used in Performance Test 2 (described in Example 1), except that the feed for Performance Test 6 consisted of 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% normal butane composition.

Table 5 lists the online gas chromatography analysis results of the total product stream of performance tests 5, 6 and 3. From the compositional data obtained from gas chromatographic analysis, the initial ethane, propane, n-butane and total conversions were calculated according to the following formula:

Ethane conversion, %=100×(wt% ethane in the feed-wt% ethane in the output stream)/(wt% ethane in the feed)

Propane conversion, %=100×(wt% propane in feed-wt% propane in output stream)/(wt% propane in feed)

Conversion rate of n-butane, %=100×(wt% n-butane in feed-wt% n-butane in output stream)/(wt% n-butane in feed)

Total conversion of ethane + propane + n-butane = ((wt% of ethane in the feed × ethane conversion %) + (wt% of propane in the feed × conversion of propane %) + (wt% of n-butane in the feed %×n-butane conversion %))/100

table 5

For performance test 6, the negative values for % ethane conversion reported in Table 5 indicate that the amount of ethane produced as a by-product of propane and/or butane conversion exceeds the amount of ethane converted in this test. However, as can be seen from Table 5, the one-stage process produces 49.57 wt% total aromatics from a given ethane/propane/n-butane feedstock, while the two-stage process is based on 100 wt% total feed into Stage 1 unconverted ethane was fed to stage 2, resulting in 53.20 wt% total aromatics. In true two-stage operation, it is also possible that the feed to stage 2 includes all non-aromatic hydrocarbons in the output of stage 1 except for the fuel gas (methane and hydrogen). These non-aromatics would include not only ethane but also ethylene, propylene, propane, etc. based on 100 wt% total feed to stage 1, which would potentially increase the total aromatics yield to slightly over 54 wt%.

Example 6

Method configuration comparison

6.1 One-stage method (comparative)

Figure 1 is a schematic flow diagram illustrating the production of aromatics (benzene and higher aromatics) from a feed containing 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% butane using a reactor-regenerator one-stage process diagram of the method.

25 tons/hr (tph) mixed feed (stream 1) mainly composed of 31.6wt% ethane, 29.5wt% propane and 38.9wt% butane (including a small amount of methane, butane, etc.), and mainly composed of ethane Mixed with recycle stream 2 consisting of other hydrocarbons which may include but not limited to ethylene, propylene, methane, butane and some hydrogen. At this point the total feed stream 3 is introduced to the single stage aromatization reactor 100 . The aromatization reactor system can be a fluidized bed, moving bed or circulating fixed bed design. Here a circulating fixed bed design is used. The reactor system used "Catalyst A" as described in Example 5 above. Unconverted reactants as well as products exit reactor 100 via stream 4 and are fed to the separation system. Unconverted reactants and light hydrocarbons are recycled back to reactor 100 in stream 2, while the separation system yields fuel gases (mainly methane and hydrogen in stream 8 from vapor-liquid separator 200), C9+ liquid products and benzene, toluene, and xylenes (BTX).

Reactor 100 operates at about 1 atmosphere of pressure and at a temperature of 675° C., while regenerator 300 operates at around 730° C., in which coke formed in reactor 100 is removed. The heat required for the reaction step is provided by the preheated hot catalyst solids mixture during the regeneration step (9). In the regeneration step, the coke-containing catalyst flows through stream 5 to regenerator 300 and is supplied with stripping gas. The regenerated catalyst flows back to the reactor 100 via stream 6 and the stripping gas exits the regenerator 300 via stream 7 .

Reactor 100 achieved nearly complete conversion (greater than 99%) of propane and butane. The average conversion rate of single-pass mixed feed is 78.84%. As shown in Figure 1, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 6 below. This one-stage mode of operation produces about 9.7 tph of benzene (from column 400 via stream 10), 3.8 tph of toluene (from column 500 via stream 11) and 0.6 tph of mixed xylenes (from column 500 via stream 12). 600), resulting in a total BTX yield of 56.2 wt% and a total liquid yield of 67.9 wt%, relative to the mixed feed. Fuel gas generation (stream 8) was 8 tph which was 31.9 wt% of the mixed feed.

6.2 Two-stage process

Figure 2 is a schematic flow diagram for the production of aromatics (benzene and higher aromatics) from a feed containing 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% butane using the two-stage reactor-regenerator system of the present invention. ).

A mixed feed (stream 1) of 25 tons/hr (tph), mainly composed of 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% butane and including a small amount of methane, butane, etc. (stream 1), was Feed to Stage 1 aromatization reactor 100 using "Catalyst A" described in Example 1. The first stage reactor 100 operates at about 1 atmosphere pressure and at a temperature of about 600°C, while the stage 1 regenerator 200, which removes the coke formed in the reactor 100, operates at about 730°C. The heat required for the reaction step is provided by the hot catalyst solids mixture which is preheated during the regeneration step. As shown in Table 5 for Performance Test 6, Reactor 100 achieved almost complete conversion of butane and 98% conversion of propane, with negative % ethane conversion recorded. This indicates that the amount of ethane produced as a by-product of propane and/or butane conversion exceeds the amount of ethane converted in this test. Reactor effluent stream 3a is then mixed with the reactor effluent (stream 3b) of the second stage reactor 300 described below. The combined effluent from the two stage reactors (stream 4) is then fed to a separation system where the unconverted reactants consist primarily of ethane and some other hydrocarbons which may include ethylene, propane, propylene, methane, butane and light hydrocarbons, along with some hydrogen, are used as the feed (stream 2) to stage-2 aromatization reactor 300, which uses "Catalyst A" described above.

The second stage reactor 300 operates at about 1 atmosphere and a temperature of about 620°C, while the regenerator 400 operates at about 730°C, which removes coke formed in the reactor. The heat required for the reaction step is provided by the preheated hot catalyst solids mixture during the regeneration step. The second stage reactor 300 converted almost half of the ethane fed to it, as was the case for Performance Test 4 in Table 3 above. The effluent from the second stage reactor 300 (stream 3b) is mixed with the effluent from the first stage reactor 100 described above. Both Stage-1 and Stage-2 of the aromatization reactor system use a circulating fixed bed design.

The average single pass conversion of the mixed feed was obtained from the cumulative conversions of ethane, propane and butane (feed) in the two stages and was calculated to be 74.5%. As shown in Figure 2, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 6 below. This two-stage operation produces about 8.9 tph of benzene (from column 600 via stream 10), 5.6 tph of toluene (from column 700 via stream 11), and 1.2 tph of xylenes (from column 800 via stream 12), compared to From the mixed feed, the resulting total BTX yield was 62.6 wt% and the total liquid yield was 72.4 wt%. Undesired fuel gas generation (stream 8 from vapor-liquid separator 500) was about 6.8 tph, which was about 27.3 wt% of the mixed feed.

6.3 Comparison of Process Configurations

Table 6 below shows a comparison of the system performance of the one-stage and two-stage methods. Conditions under which the process yielded a constant overall feed conversion were compared. Compared to each of the two stages in a two-stage process, a single-stage process must be operated at higher temperatures to achieve similar feed conversions per pass. It is evident from Table 6 that the two-stage operation produced higher product yields of benzene, toluene, mixed xylenes and C9+ liquids with less undesirable fuel gas formation compared to the one-stage process.

Table 6

Notice:

• All yields are expressed as tons of product per ton of mixed feed entering the overall process, expressed as a percentage.

The average conversion rate per pass of the two-stage process is calculated by the following formula:

(total propane conversion x mole fraction of propane in the mixed feed) + (total butane conversion x mole fraction of butane in the mixed feed) + (total ethane conversion x mole fraction of ethane in the mixed feed )

Example 7

In this example, the results of laboratory tests are used to illustrate a one-stage aromatization process compared to a two-stage process utilizing the same catalyst in each stage. The lower alkane raw material in this embodiment is composed of 31.6wt% ethane, 29.5wt% propane and 38.9wt% n-butane, and the temperature of the second stage is the same as that of the first stage.

A fresh 15-cc charge of Catalyst A (prepared as described above in Example 1) was subjected to performance testing as described below. Performance Test 5 (described in Example 5) was conducted under conditions for a one-stage aromatization process using a mixed ethane/propane feed. Performance Test 7 was conducted at the conditions used in the first stage of a two-stage aromatization process of the present invention using a mixed ethane/propane/butane feed. Performance Test 3 (described in Example 1) was performed under the conditions used for the second stage of the two-stage aromatization process of the present invention.

Performance test 7 was carried out in the same manner and under the same conditions as used for performance test 6 (described in Example 5), except that performance test 7 used a reactor wall temperature of 621°C.

Table 7 lists the online gas chromatography analysis results of the total product stream of performance tests 5, 7 and 3. Based on the compositional data obtained from gas chromatographic analysis, the initial ethane, propane, n-butane and total conversions were calculated according to the formula given in Example 5.

Table 7

Performance test 5 7 3 Catalyst A A A

Catalyst volume, cc 15 15 15 Reactor wall temperature, ℃ 675 621 621 Pressure, MPa 0.1 0.1 0.1 Feed composition Ethane, wt% 31.6 31.6 100 Propane, wt% 29.5 29.5 -0- n-Butane, wt% 38.9 38.9 -0- Gross Feed Rate, GHSV 1000 1000 1000 Total Feed Rate, WHSV 2.24 2.24 1.61 Ethane conversion rate, % 34.02 -9.61 49.28 Propane conversion rate, % 99.27 98.35 -  N-butane conversion rate, % 99.80 99.79 - Total conversion rate of ethane+propane+n-butane, % 78.84 64.77 49.28 Reactor discharge composition, wt% hydrogen 4.93 4.41 4.71 methane 18.2 11.71 7.56 Vinyl 5.57 3.61 3.95 Ethane 20.86 34.65 50.72 Propylene 0.57 0.52 0.58 propane 0.22 0.49 0.70 C4 0.08 0.08 0.11 C5 0 0 -0- Benzene 28.33 20 16.60 Toluene 11.1 12.21 8.72 C8 aromatic hydrocarbons 1.62 2.45 1.70 C9+ aromatic hydrocarbons 8.52 9.87 4.65 Total Aromatic Hydrocarbons 49.57 44.53 31.67

The negative % ethane conversion reported in Table 7 for Performance Test 7 indicates that the amount of ethane produced as a by-product of propane and/or butane conversion exceeds the amount of ethane converted in this test. However, as can be seen from Table 7, the one-stage process produces 49.57 wt% total aromatics from a given ethane/propane/n-butane feedstock, while the two-stage process schedule is based on 100 wt% total feed into stage 1, followed by Unconverted ethane in stage 1 was fed to stage 2, resulting in 55.50 wt% total aromatics. In true two-stage operation, it is also possible that the feed to stage 2 includes all non-aromatic hydrocarbons in the output of stage 1 except for the fuel gas (methane and hydrogen). These non-aromatics would include not only unconverted ethane but also ethylene, propylene, propane, etc. based on 100 wt% total feed to stage 1, which would make it possible to increase the total aromatics yield to slightly over 56 wt %.

Example 8

Method configuration comparison

8.1 One-stage method (comparative)

Figure 1 is a schematic flow diagram illustrating the production of aromatics (benzene and higher aromatics) from a feed containing 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% butane using a reactor-regenerator one-stage process diagram of the method.

25 tons/hr (tph) mixed feed (stream 1) mainly composed of 31.6wt% ethane, 29.5wt% propane and 38.9wt% butane (including a small amount of methane, butane, etc.), and mainly composed of ethane It is mixed with a recycle stream 2 consisting of other hydrocarbons which may include but are not limited to ethylene, propane, propylene, methane, butane, and some hydrogen. At this point the total feed stream 3 is introduced to the single stage aromatization reactor 100 . The aromatization reactor system can be a fluidized bed, moving bed or circulating fixed bed design. Here a circulating fixed bed design is used. The reactor system used "Catalyst A" as described in Example 7 above. Unconverted reactants as well as products exit reactor 100 via stream 4 and are fed to the separation system. Unconverted reactants and light hydrocarbons are recycled back to reactor 100 in stream 2, while the separation system yields fuel gases (mainly methane and hydrogen in stream 8 from vapor-liquid separator 200), C9+ liquid products and benzene, toluene, and xylenes (BTX).

Reactor 100 operates at about 1 atmosphere of pressure and at a temperature of 675° C., while regenerator 300 operates at around 730° C., in which coke formed in reactor 100 is removed. The heat required for the reaction step is provided by the preheated hot catalyst solids mixture during the regeneration step (9). In the regeneration step, the coke-containing catalyst flows through stream 5 to regenerator 300 and is supplied with stripping gas. The regenerated catalyst flows back to the reactor 100 via stream 6 and the stripping gas exits the regenerator 300 via stream 7 .

Reactor 100 achieved nearly complete conversion (greater than 99%) of butane and propane. The average conversion rate of single-pass mixed feed is 78.84%. As shown in Figure 1, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 6 below. This one-stage mode of operation produces about 9.7 tph of benzene (from column 400 via stream 10), 3.8 tph of toluene (from column 500 via stream 11) and 0.6 tph of mixed xylenes (from column 500 via stream 12). 600), resulting in a total BTX yield of 56.2 wt% and a total liquid yield of 67.9 wt%, relative to the mixed feed. Fuel gas generation (stream 8) was 8 tph which was 31.9 wt% of the mixed feed.

8.2 Two-stage process

Figure 2 is a schematic flow diagram for the production of aromatics (benzene and higher aromatics) from a feed containing 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% butane using the two-stage reactor-regenerator system of the present invention. ).

25 tons/hr (tph) of mixed feed (stream 1), which is mainly composed of 31.6 wt% ethane, 29.5 wt% propane and 38.9 wt% butane including a small amount of methane, butane, etc. (stream 1), was Feed to Stage 1 aromatization reactor 100 using "Catalyst A" described in Example 1. The first stage reactor 100 operates at about 1 atmosphere pressure and at a temperature of about 620°C, while the stage 1 regenerator 200, which removes the coke formed in the reactor 100, operates at about 730°C. The heat required for the reaction step is provided by the preheated hot catalyst solids mixture during the regeneration step. As shown in Table 7 for Performance Test 7, Reactor 100 achieved almost complete conversion of butane and propane, with negative % ethane conversion recorded. This indicates that the amount of ethane produced as a by-product of propane and/or butane conversion exceeds the amount of ethane converted in this test. Reactor effluent stream 3a is then mixed with the reactor effluent (stream 3b) of the second stage reactor 300 described below. The combined effluent from the two stage reactors (stream 4) is then fed to a separation system where the unconverted reactants consist primarily of ethane and some other hydrocarbons which may include ethylene, propane, propylene, methane, butane and light hydrocarbons, along with some hydrogen, are used as the feed (stream 2) to stage-2 aromatization reactor 300, which uses "Catalyst A" described above.

The second stage reactor 300 operates at about 1 atmosphere and the same temperature (620°C) as the first stage, while the regenerator 400 operates at around 730°C, which removes the coke formed in said reactor. The heat required for the reaction step is provided by the hot catalyst solids mixture which is preheated during the regeneration step. The second stage reactor 300 converted almost half of the ethane fed to it, as was the case for Performance Test 3 in Table 7 above. The effluent from the second stage reactor 300 (stream 3b) is mixed with the effluent from the first stage reactor 100 described above. Both Stage-1 and Stage-2 of the aromatization reactor system use a circulating fixed bed design.

The average single pass conversion of the mixed feed was obtained from the cumulative conversions of ethane, propane and butane (feed) in the two stages and was calculated to be 80.9%. As shown in Figure 2, the liquid product is separated in a series of three columns in series to obtain separated liquid products. The yields of the process are summarized in Table 8 below. This two-stage operation produces approximately 8.6 tph of benzene (from column 600 via stream 10), 5 tph of toluene (from column 700 via stream 11), and 1 tph of xylene (from column 800 via stream 12), relative to the combined Feed, resulting in a total BTX yield of 58.5 wt% and a total liquid yield of 72.5 wt%. Undesired fuel gas generation (stream 8 from vapor-liquid separator 500) was about 6.8 tph, which was about 27.2 wt% of the mixed feed.

8.3 Comparison of Method Configurations

Table 6 below shows a comparison of the system performance of the one-stage and two-stage approaches. Conditions under which the process yielded a constant overall feed conversion were compared. Compared to each of the two stages in a two-stage process, a single-stage process must be operated at higher temperatures to achieve similar feed conversions per pass. It is evident from Table 8 that the two-stage operation produced higher product yields of benzene, toluene, mixed xylenes and C9+ liquids with less undesired fuel gas formation compared to the one-stage process.

Table 8

Notice:

• All yields are expressed as tons of product per ton of mixed feed entering the overall process, expressed as a percentage.

The average conversion rate per pass of the two-stage method is calculated as follows:

(total propane conversion x mole fraction of propane in the mixed feed) + (total butane conversion x mole fraction of butane in the mixed feed) + (total ethane conversion x mole fraction of ethane in the mixed feed )

Claims (15)

1. A method for converting mixed lower alkanes into aromatic hydrocarbons, comprising: firstly converting a mixed lower alkane feed comprising at least propane and ethane in the presence of an aromatization catalyst to convert the first stage of aromatic hydrocarbons to maximum propane reaction under the first stage reaction conditions of aromatic reaction products, ethane is separated from said first aromatic reaction products, and ethane is converted into second stage aromatic reaction products in the presence of an aromatization catalyst to maximize the conversion of ethane reacting under the second stage reaction conditions and optionally separating the ethane from the second stage aromatic reaction product. 2. The method of claim 1, wherein the aromatization reaction is carried out at a temperature of 400 to 700°C. 3. Process according to claims 1 and 2, wherein the first stage reaction conditions comprise a temperature of 400 to 650°C, preferably a temperature of 420 to 650°C. 4. Process according to claims 1 to 3, wherein the second stage reaction conditions comprise a temperature of 450 to 680°C, preferably a temperature of 450 to 660°C. 5. A process as claimed in claims 1 to 4, wherein the first stage reaction product is produced in at least two parallel reactors. 6. A process as claimed in claims 1 to 5, wherein the second stage reaction product is produced in at least two parallel reactors. 7. A process as claimed in claims 1 to 6 wherein other non-aromatic hydrocarbons are produced in the first stage and reacted with ethane in the second stage to produce additional second stage aromatic reaction products. 8. A process as claimed in claims 1 to 7 wherein a fuel gas is also produced in either or both of the first and second stages and is separated from the aromatic reaction products and ethane. 9. A method for converting mixed lower alkanes into aromatic hydrocarbons, comprising: firstly feeding mixed lower alkanes comprising at least propane and ethane in the presence of an aromatization catalyst in a process that maximizes propane and said feed conversion of any other higher hydrocarbons present to a first stage aromatic reaction product under first stage reaction conditions, combining said first aromatic reaction product with ethane and any other non- Aromatics separation, converting ethane and at least a portion of any other non-aromatic hydrocarbons produced in said first stage in the presence of an aromatization catalyst to maximize conversion of ethane and said other non-aromatic hydrocarbons to second-stage aromatics reacting under the second stage reaction conditions of aromatic reaction products, and optionally separating said second stage aromatic reaction products from ethane and other non-aromatic hydrocarbons. 10. The method of claim 9, wherein the aromatization reaction is carried out at a temperature of 400 to 700°C. 11. A process as claimed in claims 9 and 10, wherein the first stage reaction conditions comprise a temperature of 400 to 650°C, preferably a temperature of 420 to 650°C. 12. Process according to claims 9 to 11, wherein the second stage reaction conditions comprise a temperature of 450 to 680°C, preferably a temperature of 450 to 660°C. 13. A process as claimed in claims 9 to 12, wherein the first stage reaction product is produced in at least two parallel reactors. 14. A process as claimed in claims 9 to 13 wherein the second stage reaction product is produced in at least two parallel reactors. 15. A process as claimed in claims 9 to 15 wherein fuel gas is also produced in either or both of the first and second stages and is separated from the aromatic reaction products and ethane.

Images